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Design of Sour Water Stripping System
Conference Paper · February 2009
DOI: 10.13140/RG.2.1.3663.0801
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Design of Sour Water Stripping System
Jed M. Bellen
Fluor Daniel Inc. Philippines, 3rd Floor Asian Star Building
2402 – 2404 ASEAN Drive, Filinvest Corporate City, Alabang
Muntinlupa City 1781 Philippines
Telephone Number: (632) 8504451
Email: Jed.Bellen@Fluor.com
Abstract:
Among the vital units in a refinery is the Sulfur Recovery Unit, where
elemental sulphur is recovered from gaseous hydrogen sulphide. Part of the Sulfur
Recovery Unit is the Sour Water Stripping Section which would treat sour water
coming from various sources around the refinery.
Sour water is basically waste water which contains H2S and NH3 and
sometimes phenol which can be treated by stripping in a Sour Water Stripper to
remove the H2S and NH3 content (or phenol) so that the water can be reused, further
treated or discarded to sewerage if it met the required quality.
In this paper, the general approach to the design of sour water stripping section
is described which involves process simulation, design and operating conditions map
preparation, material selection diagram preparation, flash drum, feed tank and stripper
sizing, tray and heat exchanger rating, hydraulic calculations and pump design.
As sour water treatment is very important in every refinery, it is therefore
desirable to come up with an optimal design of sour water stripping systems.
Key words: Sulfur Recovery Unit, Sour Water, Sour Water Stripper, Refinery,
Optimal Design
Introduction
Among the most vital units in a refinery is the Sulfur Recovery Unit where
elemental sulphur is recovered from gaseous hydrogen sulphide. Part of the Sulfur
Recovery Unit is the Sour Water Stripping Section which would treat sour water
coming from various sources around the refinery.
Sour water is produced in Atmospheric Crude Columns and Vacuum Crude
Towers when stripping steam is condensed and removed by overhead condensing
systems. It is also produced in Vacuum Crude Towers from equipments such as
ejectors and barometric condensers which are designed to maintain vacuum inside the
column. Steam injection to vacuum heater is another source of sour water in Vacuum
Distillation Unit. In Thermal and Catalytic Cracking Units, sour water is produced as
condensates from steam used in injection, stripping and aeration. Hydrotreater wash
water is also another major source of sour water.
Heavy viscous feeds which are rich in sulphur produce high H2S concentrations
when hydrogenated while ammonia (NH3) is produced from hydrogenation of organic
nitrogen compounds. If more sulphur will be removed to meet the more stringent
environmental requirements then there may be more nitrogen converted to ammonia
which would accumulate in wash water. Hydrogen sulphide and ammonia
concentrations are highest in sour water coming from Hydrodesulfurization (HDS)
Units and Fluid Catalytic Cracking (FCC) Units. In addition to this, phenols are
produced from reactions between steam and cyclic hydrocarbons. Sour water with
extremely high concentration of phenols would come from the FCC Units.
Sour water systems are garbage disposals or toilets of refineries. It does not
receive constant feed rate and composition. Any water soluble waste produced in the
refinery, either continuously, intermittently or in slug will be disposed into this system
(Armstrong, et al., 1996).
Process Description
The sour water from various sources in the refinery is coursed to the Sour
Water Stripping Unit to strip the water of its sour content like H2S, NH3 and phenol.
The sour water is first fed to a Flash Drum where hydrocarbon vapours and
liquids are removed. The vapours are flashed through a pressure controlled vent
connected to a low pressure system like sour water system flare knock out drum. The
collected hydrocarbons are pumped to a slop system where it could be reprocessed.
The sour water from the flash drum is then fed to the Feed Stabilization Tank.
This stabilization tank is used to increase the residence time of the sour water for
longer mixing and homogenization of the feed composition and for further removal of
hydrocarbons. If this is not done and it happens that sour water composition changes
significantly, the stripper will not operate properly and it could result to inconsistency
in product specification or steam wastage by overstripping leading to the worst
composition. Furthermore, environmental regulations might require vent gases to be
treated for hydrocarbon or H2S (Armstrong et al., 1996).
The sour water from the tank is sent to a heat exchanger where it is pre-heated
through heat exchange with the stripped water. Then it flows to the Sour Water
Stripper where some of it is flashed through a reduction in the pressure across a
control valve. There are varieties of sour water stripping methods employed. Most of
them involve the flow of sour water through trays or packings in a column while the
stripping steam or gas flowing upwards removes H2S and sometimes NH3 (Beychock,
1967). The steam that could be used maybe live steam or steam produced from the
reboiler. The reboiler boils sour water at the minimum tower operating pressure by
utilizing the latent heat of low pressure steam as the heating medium. This has the
advantage of no additional water load in the stripper and the steam condensate can be
recovered and returned to the boiler house. The H2S, NH3 and steam rising to the
tower cooling section are cooled by pumped-around sour water in the middle of the
tower. The temperature is maintained at 1800F. The reason for this is that if the
temperature would be well above this value, there could be carry over of condensates
into the overhead due to high vapour flow. On the other hand, if the temperature is
well below this value there could be no sufficient removal of H2S in the sour water.
The overhead gases flow by pressure control to a lower pressure system like the Sour
Water Gas Flare Knock Out Drum or Sulphur Recovery Unit. The stripped water
collected at the tower bottom flows through a heat exchanger where it is cooled via
heat exchange with sour water feed to the tower. It is then pumped off-site on level
control for further processing. It could be sent to a crude unit desalter; a liquid/liquid
extractor that transfer ninety-five percent (95%) of phenols in water to atmospheric
crude feed. Phenol removal in sour water stripper is minimal (~ 10% reduction) but if
it is sent to the desalter, phenol content is reduced substantially.
The Process Flow Diagram is shown in Figure 1.
Gas to
Flare K.O.
Drum
Vapor to
Flare K.O.
Drum
Gas to Sulfur
Recovery Unit
Drum
PC
Sour Water
Feed
TIC
PIC
LIC
FLASH
DRUM
SOUR WATER
STRIPPING
TOWER
Hydrocarbon
Liquids
LIC
FIC
FIC
LP
STEAM
COND.
FEED
STABILIZATION
TANK
FIC
STRIPPED
WATER TO
DESALTER
Figure 1. The Process Flow Diagram of Sour Water
Stripping Systems
Sour Water Chemistry
Sour water is an aqueous solution of NH3 and H2S which may contain as much
as 10,000 ppm of H2S. The NH3 to H2S ratio ranges from 1 to 2 usually 1.5 as the
average. The alkalinity ranges from pH 7.8 to pH 9.3. The NH3 and H2S are present in
aqueous solution as NH4SH which is the salt of a weak base (NH4OH) and a weak
acid (H2S).
This salt would undergo hydrolysis in solution to form back H2S and
NH3. The hydrolysis reaction could be represented by Figure 2.
(Vapor Phase)
(Aqueous Phase)
NH3
P.P.
H2S
P.P.
NH4++SH- < - - > NH3 + H2S
Figure 2. Overall Equilibrium in Sour Water
For a detailed discussion on sour water chemistry, Ref. 1 can be consulted. It provides
a discussion on the pH of sour water and vapour pressure of NH3 and H2S above
aqueous NH4SH solutions. It also provides useful graphs showing the correlation of
concentration of NH3 (in ppm) in aqueous solution versus the pH of the solution;
temperature versus correlation factor for partial pressure of H2S over aqueous
solutions of NH3 and H2S; and Henry’s Law coefficient for pure water versus the
temperature and vapour pressures of H2S and NH3 over the aqueous solutions of H2S
and NH3 at the temperatures of 200 0F, 210 0F, 220 0F, 225 0F and 230 0F .These
graphs were based on the study conducted by Van Krevelin, et al.
Process Simulation
Recent simulation software like HYSYS or PRO/II has “drag and drop”
features that make simulation relatively easy. Other software like Aspen Plus,
ProMax, TSWEET and PROSIM can also be used. Simulating Sour Water Stripping
Systems would have direct convergence since both mass and heat can flow in and out
of the system.
Since the sour water that would be fed into the flash drum is at saturated
conditions; then, in the flash drum, it is expected that there would be liquid-vapour
separation, though the flashed vapour may be minimal. Another flash drum could
represent the feed stabilization tank but, in here, it is expected that liquid water comes
out of the vessel. The sour water would then pass through the Feed-Bottoms Heat
Exchanger to preheat the sour water to 180 0F by heat exchange with the stripped sour
water. The preheated Feed Sour water is fed to the Sour Water Stripper. The Feed
Tray and the Number of Trays are determined through adjustments so that the
Stripped Water specified quality will be met. The number of trays that is obtained here
is the theoretical number of trays. Since the stripper has pumparound, the steam going
in and out of the condenser can be represented by pseudo-streams for tractable stream
flow rates and properties. If a kettle reboiler is used, the fraction of liquid to be
vaporized needs to be specified. The H2S vapour coming out of the stripper could be
sent to the Sulfur Recovery Unit to recover the sulfur. Finally, the stripped sour water
coming out of the Stripper bottom is cooled through heat exchange with the feed.
In PRO/II there are two (2) methods for use in modelling the system
equilibrium. These are SOUR and GPSWAT methods. The SOUR method is based on
Sour Water Equilibrium (SWEQ) model developed by Grant Wilson for the American
Petroleum Institute (API) and the Environmental Protection Agency (EPA). This
method contains four components namely H2O, NH3, H2S and CO2. The correlation is
based on the components partial pressures. Similarly, the GPSWAT method was
developed by Grant Wilson for the GPA. This method contains nineteen components,
including H2O, NH3, H2S and CO2. This is a rigorous model of the reactive equilibria
in the system.
Both these methods can be used for rigorous Vapor-Liquid-Liquid Equilibrium
(VLLE) calculations and can be used with petroleum fractions. The keyword input for
thermodynamic data is provided in the Appendix 1 (Lecture of Dr. Jungho Cho).
See Figure 3 for the sample HYSYS Simulation of Sour Water Stripping
System.
Figure 3. HYSYS Simulation of Sour Water Stripping Systems
Determination of Design and Operating Conditions
The design and operating conditions of the flash drum, feed tank, sour water
stripper and internals, heat exchangers and pumps need to be specified to help the
equipment engineer determine the equipment’s thickness and all the other
requirements specified in the applicable code. Also the pipeline design and operating
conditions need to be determined to help the piping engineer in computing for the pipe
thickness and other various requirements by applicable code. See Figure 4 for the
Design and Operating Conditions Map.
Legend:
DP = Design Pressure
DT = Design Temperature
OP = Operating Pressure
OT = Operating Temperature
Vapour to
Flare K.O.
Drum
DP= Design Battery Limit Pressure
DT= OT + 500F
OP= Normal Battery Limit Pressure
OT= Feed Temp.
Sour Water
Feed
PIC
FLASH
DRUM
LIC
DP= From Standard
DT= OT +500F
OP= Saturation
Pressure
OT= Feed Temp.
TIC
SOUR WATER
STRIPPING
TOWER
DP= Pump Shut
off/Full Vacuum
DT= OT + 500F
OP= ATM
OT= Feed Temp.
LIC
FIC
LP
STEAM
DP= From Std.
DT= 250 0F
OP= 10 psig
OT= 200 0F
FIC
DP= Pump Shut-off
DT= OT + 500F
OP= Pump
Discharge minus DP
across the Control
Valve
OT= Feed Temp.
FIC
COND.
FEED
STABILIZATION
TANK
DP= Pump Shut-off
DT= OT + 500F
OP= Pump Discharge
OT= Feed Temp.
Gas to Sulfur
Recovery Unit
Drum
PC
DP= From Standard
DT= OT + 50 0F
OP= From Simulation
OT= From Simulation
DP= Pump Shut-off
DT=230 0F
OP= Heat Ex. Outlet
Pressure
OT= 180 0F
Hydrocarbon
Liquids
DP= Max. Suction
DT= OT + 500F
OP= Pump
Suction
OT= Feed Temp.
Gas to
Flare K.O.
Drum
DP= Max. Suction
DT= 2500F
OP= Shell Inlet
Pressure
OT= 2000F
DP= Pump Shut-off (In &
Out)
DT= OT + 500F (In & Out)
OP= Tube Inlet Pressure (In)
= Inlet Pressure – ΔP (Out)
OT= Feed Temp (In)
= Outlet Temp (Out)
Figure 4. Design and Operating Conditions Map
Material Selection
Basically, the material that can be used for the vessels, tanks, stripper, pumps,
heat exchangers and piping is carbon steel. Higher grade killed carbon steel can also
be used. See the Material Selection Diagram depicted in Figure 5.
STRIPPED
WATER TO
DESALTER
Vapor to
Flare K.O.
Drum
Gas to
Flare K.O.
Drum
Mat’l: CS
CA: 1/8 in
Gas to Sulfur
Recovery Unit
Drum
PC
Sour Water
Feed
Mat’l: CS with Ti
TIC
Mat’l: CS
CA: 1/8 in
Mat’l: CS
CA: 1/8 in
PIC
FLASH
DRUM
LIC
Mat’l: CS
CA: 1/8 in
Mat’l: CS
CA: 1/8 in
Hydrocarbon
Liquids
SOUR WATER
STRIPPING
TOWER
Mat’l: CS
CA: 1/8 in
LIC
FIC
FIC
LP
STEAM
COND.
FEED
STABILIZATION
TANK
FIC
Mat’l: CS
CA: 1/8 in
Mat’l: CS
CA: 1/8 in
Mat’l: CS
CA: 1/8 in
Mat’l: CS
CA: 1/8 in
Legend: Mat’l = Material
CA = Corrosion Allowance
Figure 5. Material Selection Diagram
Equipment Design
Flash Drum
The sour water introduced to the drum contains minimal amount of oil which
should be removed by flotation to protect the downstream equipment from fouling and
foaming in the stripping column. Knowing that the oil is less dense than water, it
would float above the water; and to achieve separation, a weir could be installed near
the other end of the drum opposite the feed inlet. The oil would overflow from this
weir and would be collected to the smaller chamber. See Figure 6. In this
configuration, level is controlled to about three (3) inches from the top of the weir.
Mat’l: CS
CA: 1/8 in
Stripped
Water to
Desalter
Figure 6. Flash Drum with Overflow Weir
Another way of collecting the oil is to install inside the flash drum a draw-off
box which would collect the oil overflowing into it. See Figure 7.
Figure 7. Flash Drum with Draw-off Box
Drawing off the oil is done when the chamber or the draw-off box is already
full. The oil is then sent to the designated refinery sump. The designated residence
time for the sour water inside the drum is usually twenty (20) minutes.
The flash drum can be sized using the following steps as provided by Branan
(2005):
Step 1. Calculate the separation factor, S.F.
0.5
W ⎛ρ ⎞
S.F. = L ⎜⎜ V ⎟⎟
WV ⎝ ρL ⎠
where: W = liquid or vapour rate
ρ = density of liquid or vapor
Step 2. Look up for system constant, KH. The Figure is provided by Branan (2005) in
Ref. 1, Figure 1, page 132.
Step 3. Calculate the maximum vapour velocity, Uvapour max.
U vapor max
⎡ (ρ − ρ V ) ⎤
= KH ⎢ L
⎥
⎣ ρV
⎦
0. 5
Step 4. Calculate the required vapour flow area.
AV min =
QV
U vapor max
where: QV = vapour volumetric flow rate
AVmin = required vapour flow area
Step 5. Select the appropriate design surge time and calculate the full liquid volume.
Step 6. When vessel is at full liquid volume:
Atotalmin =
AV min
0.2
Dmin = (4[Atotal ]min / π )
0.5
where: Dmin = minimum diameter
Step 7. Calculate the vessel length.
L=
FLV
⎛π ⎞ 2
⎜ ⎟D
⎝4⎠
where: FLV = Full Liquid Volume
D = Dmin to the next largest 6 in
Step 8. If 5 < L/D < 3, resize.
Feed Stabilization Tank Sizing
The sour water from the flash drum is directed to the feed tank. This tank could
be a fixed or a floating-roof type. It could be operated to about 60% full. About two
(2) feet of hydrocarbons should be allowed to float above the sour water to reduce
odour. For cone roof tanks, nitrogen blanketing is required to control odour and to
prevent the formation of explosive mixtures. The sour water is pumped on flow
control to the Sour Water Stripper.
The tank is sized using the volume equation for cylinder if the volume to be
contained is known. The very important volume to consider is the tank’s working
volume.
Sour Water Stripper Design
There are two types of sour water stripper: a steam stripper and a reboiled
stripper. The design configurations of these two are discussed in the forgoing
paragraphs.
Steam Strippers
In this stripper, the stripping media used is steam. The operating condition
varies from 1 to 50 psig and from 100 to 270 0F. The sour water may or may not be
pre-treated with mineral acid like H2SO4 or HCl before stripping (Beychock, 1967).
A major portion of stripping steam is used in heating the sour water feed from
180 0F to 230 0F. The heat that is being utilized is the latent heat of the steam. A
portion of the stripping steam breaks the bond between H2O and H2S and the bond
between H2O and NH3. When these species dissolves in water, it evolves heat called
the heat of solution. In order for those components to be removed from water, the
same amount of heat must be supplied to the solution. The heat that is supplied comes
from the latent heat of steam condensing across the trays in the tower. Some of the
stripping steam condenses in the overhead condenser. The condensed steam which
accumulates in the reflux drum is totally refluxed back to the top tray of the tower.
The small amount of stripping steam remaining as vapour leaves the reflux drum,
together with H2S and NH3 vapour.
Using this kind of stripper, however, would add steam condensate to the
stripped water, thereby increasing the plant’s water effluent (Lieberman, 2003).
CW
Steam + H2S + NH3
Sour
Water
Reflux
Steam
Stripped
Water
Figure 8. Steam Stripper
Reboiled Stripper
The same amount of steam is required for this type of stripper as that of steam
stripper. The advantage of using this type of stripper is that the steam condensate can
be recovered and recycled back to boilers. However, its main disadvantage is fouling
in the reboiler wherein the causes are difficult to determine and control (Lieberman,
2003). In addition to that, reboiled strippers involve higher investment cost. Figure 9
shows the configuration of this type of stripper.
Sour Gas
TIC
Sour
Water
FIC
Steam
Stripped
Water
Figure 9. Reboiler Stripper
In sizing this stripper, tray rating is first performed using any available
software. The tray type should be specified but sieve tray is commonly used.
Basically, the size of the tray will depend on the maximum vapour and liquid load.
The tray hole diameter is specified, ranging from 1/16 to 1 inch. Also the fractional
hole areas is specified. For commercial sieve trays, the optimum value for fractional
hole area is from 0.08 to 0.12 (Kister, 1990). The downcomer width needs to be
specified and the pressure drop across the tray is closely monitored.
The height of the column would depend on tray spacing. Kister (1990)
recommended a tray space, ranging from 8 to 36 inches, though 24 inches is most
common for columns with diameter 4 feet and larger. The reason for this value is for a
space sufficiently large that a worker can crawl freely between trays.
Equally important to consider in designing a stripper column is the removal of
aromatic components in sour water because of environment requirements. To effect
the efficient removal of aromatics, saltwater is added to the stripper feed because
aromatics, like benzene, are less soluble in brine than in freshwater. However, brine is
corrosive to the stripper.
Heat Exchanger
Feed-Bottoms Exchanger
The rating of this heat exchanger is much simpler as compared to rating
reboilers. The rating could be done using software like HTRI. The most tedious part of
this activity is finding the shell diameter and tube length. However, HTRI offers the
Design Mode which can calculate the shell diameter and tube length. Upon knowing
the shell diameter and tube length, the HTRI design mode is changed to rating mode,
and those values are used as initial input for rating the heat exchanger. The
parameters, which are closely guarded, are shell side and tube side pressure drop, shell
side and tube side velocity, ρv2 and overdesign. The Overdesign is defined as the
actual overall heat transfer coefficient divided by the design overall heat transfer
coefficient. Its value must be greater than zero (0).
Reboiler
The types of reboiler which can be used are either the once-through
thermosiphon reboiler or the kettle reboiler.
Once-through Thermosiphon Reboiler
In a thermosiphon reboiler, the driving force to effect flow is the density
difference between the reboiler feed line and the froth filled reboiler return line. The
differential pressure driving force can be obtained from the values of the specific
gravity of the liquid in the reboiler feed line, the height of the liquid above the reboiler
inlet, the mixed-phase specific gravity of the froth leaving the reboiler, and the height
of the return line. The differential pressure is consumed by frictional losses in the
reboiler, the inlet line, the outlet line and the nozzles. Figure 9 shows the
configuration at the tower bottoms and another configuration is shown below.
BAFFLE
SEAL PAN
LP Steam
Stripped
Water
Figure 10. Tower Bottoms Configuration for Once-Through
Thermosiphon Reboiler
In Figure 10, all liquid from the bottom tray flows to the reboiler. No liquid from
tower bottoms flows into it. The bottoms product came from the reboilers liquid
effluent. Hence, the reboiler outlet temperature is equal to tower bottoms temperature
(Lieberman, 2003).
Kettle Reboiler
Kettle reboilers are installed external to the tower as shown in Figure 11.
LP
Steam
Stripped
Water
Figure 11. Kettle Reboiler
It is important to note that the bottoms product level control only controls the
liquid level in the product side of the reboiler and not the liquid level in the tower. The
liquid level on the boiling side of heat exchanger is controlled by the overflow weir.
The water that comes out of the tower would flow to the bottom of the reboiler
shell. Once it entered the boiling side of the heat exchanger, it is partially vaporized.
The top section of the reboiler separates the vapour and the liquid. The vapour flows
back to the tower via the riser or return line. The vapour becomes the stripping vapour
or heat source. The overflow weir height must ensure that the tube bundle is
submerged in the liquid. The liquid that flows over the weir becomes the bottoms
product (Lieberman, 2003).
Another important thing affected by the reboiler operation is the liquid level in
the tower. The liquid level in the tower is the sum of the nozzle exit loss of the liquid
leaving the bottom of the tower, the liquid feed-line pressure drop, the shell-side
exchanger pressure drop (including the effect of baffle height) and the vapour-line
riser pressure drop (including the vapour outlet nozzle loss). It is important to note,
however, that it is the elevation and not the static head pressure in the tower that
drives the reboiler. This means that the pressure in the kettle is greater than the
pressure in the tower. Any increase in reboiler duty would correspond to an increase
in liquid level at the bottom of the tower. If the liquid level at the bottom of the tower
rise to the reboiler vapour return nozzle, the tower will flood but the duty will remain
the same. Reboiler shell-side fouling may lead to tray flooding because it can cause
pressure drop build-up on the shell side of the reboiler.
The Pumparound
There are two ways to remove heat from a distillation tower. The first one is by
introducing a top reflux and the second one is by using a circulating reflux called
pumparound. The hot liquid is drawn from the pumparound draw tray, cooled in a
condenser or air fin cooler and returned to the tower at a higher elevation. It is best
practice if this pumparound return liquid enters the tray downcomer as shown in
Figure 9.
The purpose of pumparound is to cool and partially condense the upflowing
vapors. The typical number of pumparound trays is a minimum of two (2) and a
maximum of five (5). In the case of steam stripper (See Figure 8), it is employed to
recover heat to a process stream that would otherwise be lost to the cooling water.
Hence, this could lower the cooling water outlet temperature. It is best to keep the
cooling water outlet temperature below 125 0F to retard water hardness deposition.
Pumparound is also used to prevent top-tray flooding by decreasing the vapour
temperature. In so doing, less vapour would reach the top tray resulting to lower
vapour velocity and lower tray pressure differential. In this case the ability of the
vapour to carry entrained liquid will be lessened and the height of the liquid in the
downcomer will be reduced and tray flooding will be avoided. It should be noted that
too much lowering of the vapour velocity can lead to tray weeping which is
undesirable since it does not provide good liquid-vapour contact. Also if vapour
velocity is too low, it may lead to the loss of the downcomer seal, causing vapour to
by-pass the trays via the downcomer.
The temperature difference in the pumparound trays indicates that fractionation
is taking place. The temperature difference is a measure of the amount of fractionation
as expressed in the following equation:
ΔT= temperature of liquid - temperature of vapour
leaving a lower tray
leaving a higher tray
In the equation above, the bigger the temperature difference, the more
fractionation would take place across the trays (Lieberman, 2003).
Pump Design
The type of pump commonly used in this system is centrifugal pump. When
designing a centrifugal pump the horsepower requirement needs to be determined. The
handiest formula for horsepower is
HP=GPM(ΔP)/1751(Eff)
where: HP = Pump Horsepower
GPM = Gallons per Minute
ΔP = Delivered Pressure
Eff = Pump Efficiency
Branan (2005) provided an approximate formula for pump efficiency as below:
Eff = 80 – 0.2855F + 3.78(10-4 )FG – 2.38(10-7)FG2 + 5.39(10-4)F2 –
6.39(10-7)F2G + 4(10-10 )F2G2
where: F = developed head, ft
G = Flow, GPM
The preceding equation is applicable for F=50 to 30 feet and G = 100 to 1000 GPM.
The result of the equation is within about 7% of the pump curves. For G < 100, a
rough efficiency could be obtained by using the equation for 100 GPM.
Another consideration in pump design is the provision for minimum flow.
Minimum flow is needed to protect the pump from shut-off. At shut-off, all of the
pump’s horsepower turns into heat and it can vaporize the liquid and damage the
pump. The minimum flow is constant from discharge to suction. For preliminary
estimation, assume that all the horsepower at blocked-in conditions turns into heat.
Thereafter, provide a minimum flow that could remove 15 0F rise in the minimum
flow stream temperature.
It is also essential to ensure that the fluid flowing in the pump suction line is
not vaporizing. Hence, the pressure along the line must not go below the fluid’s
vapour pressure. The lowest pressure is at the impeller inlet where a sharp pressure
lowering occurs. The impeller rapidly builds pressure and collapses the vapour
bubbles, causing cavitation and damage. To prevent this, a sufficient net positive
suction head (NPSH) must be maintained. NPSH is the pressure available at the pump
suction after vapour pressure is subtracted. In equation form:
NPSH = Absolute – Vapor – Line +
Difference
Available Pressure, ft Pressure, ft
Loss, ft
in Elevation, ft
As a rule, NPSH available must be greater than NPSH required. Maximum flow
usually has higher NPSH than normal flow. For extremely low flows, NPSH could be
higher.
In choosing the pump, an economic balance between NPSH requirements and
pump speed need to be studied. A lower speed pump would require a lower NPSH and
hence, lower vessel heights. A low speed pump will also be easy to maintain. On the
other hand, a higher speed pump delivers the required head economically.
The suction piping should be adequately sized. Normally, it is larger than pump
suction nozzle. Its layout should also be kept simple. In here, the pressure should be
kept below the vapour pressure of the fluid. To achieve this objective, Kern, as cited
by Branan (2005), provided the following rules of thumb:
1. The minimum liquid head above the drawoff nozzle should be greater than the
exit nozzle resistance. To determine the liquid level, the following formula can
be used:
hL=u2/g
where: hL = Liquid Level Above Nozzle, ft
u = Nozzle Velocity, ft/sec
g = 32.2 ft/s2
The equation is based on safety factor of 4 and a velocity head factor K of 0.5.
2. For a saturated liquid, the pipe should be vertical downward from draw-off
nozzle and must be close to the nozzle as possible. This configuration will give
the maximum static head above the horizontal portion of piping at the pump
suction. A vortex breaker should also be provided for the draw-off nozzle.
The following are guidelines for suction pipeline:
1. Layout the pipeline short and simple.
2. Avoid loop and pockets to prevent the accumulation of vapour or dirt.
3. Use an eccentric reducer with the flat side up to prevent trapping vapour as the
larger suction line transition to the pump suction nozzle.
4. The following are acceptable ΔP/100 ft:
Saturated Liquids = 0.05 to 0.5 psi/100 ft
Subcooled liquids = 0.5 to 1.0 psi/100 ft
Conclusion
This paper explores on the underlying concepts used in designing an optimal
sour water stripping systems. The sour water stripping process is described and its
simulation using PRO/II and HYSYS software is presented. It also showcased the
preparation of the Design and Operating Conditions Map and the Material Selection
Diagram. Discussions on the design considerations for sizing the flash drum, tank and
stripper and rating of trays and heat exchangers are also provided. Moreover, various
configurations for those equipments are presented. Design considerations for
pumparound and pumps are also discussed with sufficient detail.
APPENDIX 1
Keyword Input for SOUR Method:
THERMODYNAMIC DATA
METHOD KVAL (VLE) = SOUR, ENTH (L) = IDEAL, &
ENTH (V) = SRKM, DENS (L) = IDEAL, &
DENS (V) = SRKM
Keyword Input for GPSWATER Method:
THERMODYNAMIC DATA
METHOD KVAL (VLE) = GPSWAT, ENTH (L) = IDEAL, &
ENTH (V) = SRKM, DENS (L) = IDEAL, &
DENS (V) = SRKM
Acknowledgment
The author would like to thank, Fluor Daniel Inc. Philippines through the Process
Department Manager, Ma’am Corazon S. Almirez and the Former Department
Manager, Ma’am Josephine B. Sabay for their unwavering support in this endeavour. I
would also like to thank my Most Loved Mentor Ever, Ma’am Melody Lee M. Estrada
for showing me how sweet sour water can be with the enormous learning that I had
from her motherly mentoring. My gratitude also goes to Sir Raul L. Bicol, the
indomitable process engineer, for sharing to me his vast refinery experiences and for
being the reviewer of this paper. I would also like to extend my gratitude to Jay R T.
Adolacion, for reviewing and critiquing my paper. Finally, I would also like to thank
Herbie Gino S. Vinluan and Marcial N. Barroga for driving me to bring out the best
that I can be.
References
1. Beychock, Milton R. (1967) Aqueous Wastes from Petroleum and
Petrochemical Plants, England: John Wiley and Sons Ltd.
2. Branan, Carl (2005) Rules of Thumb for Chemical Engineers. 4th Ed. USA:
Elsevier.
3. Kister, Henry Z. (1990) Distillation Operation. USA: McGraw-Hill.
4. Lieberman, Norman P. And Elizabeth T. Lieberman (2003) Working Guide
to Process Equipment, 2nd Ed. USA: McGraw-Hill.
5. http://www.cbi.com/services/sour-water-stripping-units.aspx
6. http://www.cheresources.com
7. http://www.hydrocarbonprocessing.com
8. http://www.insightengineers.com/articles/SourWaterStripping.pdf
9. http://www.jaeger.com/Brochure/steamstripping.pdf
10. http://www.koch-glitsch.com/koch/product_brochures/KGIMTP.pdf
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